Start-up of platinum catalyst hydroformers



July 5, 1960 T. M. MOORE ET AL START-UP oF PLATINUM CATALYST HYDRoFoRMERs Filed Feb. l5, 1957 IIYVENTORS Thomas M. Moore Josep/1 B Goms BY l 714mg, /f Wwf g gUnited States Patent O START-UP F PLATINUM CATALYST HYDROFORMERS Thomas M. Moore, Munster, and Joseph B. Cor-ns, Grifth, Ind., assignors to Standard Oil Company, Chicago, lll., a corporation of Indiana Filed Feb. 15, 1957, Ser. No. 640,#68

Claims. (Cl. 208-65) invention relates to an improved start-up procedure for platinum catalyst hydroforming systems, and it pertains more particularly to a method vfor start-up of hydroforming systems in which platinum-alumina catalyst is contaminated with adsorbed carbon oxides.

The start-up of a platinum catalyst hydroformer requires considerable care in order to avoid catalyst deactivation. Heretofore, hydrogen has been considered a desirable start-up gas. However, when carbon oxides are present in the reaction zone or the catalyst contains adsorbed carbon oxides, such as would result from purging operations Iwith ue gas at temperatures below about 450- 500 F., hydrogen has been lfound to be unsuitable because it reduces carbon dioxide to carbon monoxide, which poisons .platinum catalyst for hydroforming, and forms water, which tends to strip promoter components from the catalyst and aggravates corrosion problems. A single purging with a hydrogen-free start-up gas, such as, bottled nitrogen or normally-gaseous hydrocarbons, removes carbon oxides in the reaction system without poisoning the catalyst, but upon heating of the catalyst bed, surprisingly large volumes of carbon oxides are desorbed from the alumina of the catalyst up to a temperature of about 450- 500 F. Repeated purges and/ or evacuations of the reactors with additional start-up gas are thus required until substantially all carbon oxides are eliminated. Such procedures are, of course, troublesome and expensive, particularly where the start-up gas must be purchased elsewhere and imported into the renery.

An object of this invention is to provide an improved start-up procedure for platinum catalyst hydroforming systems which minimizes catalyst deactivation. Another object is .to provide a procedure for starting up hydroforming systems IWith platinum-alumina catalyst containing adsorbed carbon oxides. Another object is to -provide a start-up procedure which minimizes the quantity of extraneous start-up gas required. A further object is to provide a start-up procedure which will minimize loss of promoter components from the catalyst and decrease corrosion problems. These and other objects will be apparent as the detailed description of the invention proceeds.

Platinum catalyst hydroforming systems used in commercial practice today include lead, intermediate, rand tail reactors, heating zones prior to each of these reactors, a gas separation zone after said reactors, and compressor and lines for recycling separated gas. Regenerative platinum catalyst hydroforming systems, as exemplified by Ultraforming (Petroleum Engineer, vol. XXVI, No. 4, April 1954, at page C-35), are also provided with a swing or spare reactor and regeneration facilities, including sources of ue gas and air, means for removing water from regeneration gases, a circulator, and valved lines whereby the regeneration facilities may be connected to the reaction zones when platinum catalyst requires reactivation.

We have discovered a start-up procedure for such platinum catalyst hydroformers which copes with the problem Vof carbon monoxide poisoning and doesnt require re- FPice 2 peated purging of reactors with costly start-up gas during heat-up to remove desorbed carbon oxides. Our method of start-up utilizes inexpensive and readily-available ue gas to desorb carbon oxides. Thereafter, the reaction system needs only be filled once with a start-up gas, rather than many times.

In one embodiment our method comprises the steps o purging the naphtha hydroiorming zone with ue gas, circulating ue gas containing between about 0.1 and 1.0 mol percent oxygen through said zone at a pressure in excess of about 30 pounds per square inch gage while simultaneously heating the circulating ue gas to a temperatureV above about 500 F., depressuring, introducing a start-up gas into said zone whereby the partial pressure of carbon oxides in said zone is reduced below about l pound per square inch absolute, introducing preheated naphtha charge stock at a temperature in the range of 650 to 800 F. into said zone while the pressure of the startup gas is in the range of about 30 to 200 pounds per square inch gage, continuing the introduction of naphtha charge stock at a temperature in said range until the hydrogen partial pressure in said zone is increased to at least about pounds per square inch gage, and then increasing the temperature of the naphtha charge to at least `about 850 F. Statt-up gases for purposes of this invention are nitrogen and normally-gaseous hydrocarbons, such as, methane (natural gas), ethane, propane (liquied petroleum gas), and mixtures thereof. Such gases should be substantially free of hydrogen, which, upon being introduced into the reaction zone to displace ue gas, would reduce carbon dioxide to carbon monoxide and thereby poison the catalyst. Where there has been no substantial buildup of carbon oxides in circulating ilue gas, however, and, in particular, where the reaction system is evacuated after depressuring, hydrogen may be tolerated Without substantial deactivation.

In practicing our invention, the flue gas, which is normally lthe products from combustion of hydrocarbons in air, is circulated at a pressure in excess of 30 pounds per square inch gage, e.g., 300 pounds per square inch gage, to provide sufficient suction pressure for the hydroformer recycle compressor and to provide adequate gas circulation for rapid heat-up. During flue gas circulation the carbon oxide concentration will rise with increased ternperature as more and more carbon oxides are desorbed from the catalyst. To minimize this build-up of carbon oxides in the circulating ilue gas, additional ilue gas may be introduced while portions of the circulating flue gas are vented -frorn the system. The ue gas should contain between about 0.1 to l mol percent of oxygen, preferably about 0.2 to 0.5 mol percent, to `assure substantially complete absence of carbon monoxide. Oxygen content, however, should not rise above about 1 mol percent; otherwise hazardous conditions may result, particularly when startup gas is introduced. Furthermore, in systems contaminated with metal-lic suldes, excess oxygen may lead to oxidation of the sulfur and consequent sulfation of the catalyst base.

When the catalyst bed has been heated to a temperature of about 500 F., substantially all carbon oxides are desorbed from the catalyst. The system may then be depressured to atmospheric pressure and, preferably, evacuated to a pressure below atmospheric pressure before introducing start-up gas. Such depressuring and/ or evacuation minimizes the amount of start-up gas -required to reduce the partial pressure of carbon oxides in the reaction zone to a level below about 1 pound per square inch absolute, Which level has been found from experience to cause negligible catalyst deactivation. Evacuation `is particularly desirable when start-up gas contains hydrogen to prevent any significant reduction of carbon dioxide to carbon monoxide.

After the startup Igas is Vintroduced, it should be circulated through the system to heat the catalyst to about 60G-800 F. Alternatively, the catalyst may be heated to about 60G-800 F. by circulation ofthe flue gas prior to' .depressuring evacuation, and introduction Iof the start-'up gas. Gases leaving .the tail' reactor may preferably "be, co'oledrtoV about 100'ov F;, and condensed water removed in theV gasV separation zone'. Experience has sh'oWn that temperatures of about 60G-800? F.,prefer ably 700-80097151,Y are required ,to assure. substantially complete' removalofwater 'fromfthe catalyst. The rate oifjfheating shouldjbe controlledso that the amount Vof water in the gas circulating `through Vthe catalyst beds remains-below ab'outrS .mol percent, andpreferably below Vl,inolpercent. Aqueous condensate removedin. the gas separation zone may befcorros'ive and should Vbe WithdrawnjfDuring this further heating-up andV drying step la Vpressure in excess of aboutv3`() `pounds vper` square inch gage', eg., aboutj200 pounds Aper square inchY gage, is preferred to provide suction pressure for tbe recycle compressor and adequate heat capacity.- Y t When the catalyst beds have beenheated to about i600-800 F. by the recycled gas stream and substantially lall waterhas been removed from the system, thepressure inthe system may be reduced substantially below 200 pounds per square inch gage, eg., about 3.0-150 pounds per square :inch gage. Naphtha charge may then be slowlyintroduced.V The catalyst. bed in one .of ,the reactors, preferably the swing reactor 'if the system includes one, may be raised to a somewhat higher 'temperature of `about 7'50-800 F., usually not higher than about v800" F., while the remaining reactors may he maintained at Va somewhat'lower temperature of about 650-750 F. In fa non-regenerative systemgthe initial dehydrogenation may be elected at temperatures somewhat lower than those stated hereinabove. Thus, one reactor `may be maintained at a temperature in the .range of about 650- 750513., while the remaining reactors are held at a temperature less than about 650 F., e.g.', 60Gb-650 F.

In a -system .including a swing reactor Va preferred operation is to connect the swing reactor .in tail reactor position so .that while start-.up gas is being .circulated in the rest of the system, naphtha preheated to about 750- Y800? F. is introduced directly into the vswing Vreactor and the .effluent therefrom is discharged directly through heat exchangers to the hydrogen separator. The .hydrogen thus 'formed in the swing reactor is recycled `through the Whole .system vand its concentration in .the recycled gas rapidly reaches -80 percent'or more. When the hydrogen vpartial pressure .reaches the desired level, eg., about 150 pounds persquare inch gage, or more, the incoming naphtha charge at about 800 F. transfer line temperature vis .cut into the rst reactor for on-stream flow through regular on-stream reheaters and reactors. The catalyst in .the ctn-stream reactors is thus protected by hydrogen initially generated in the swing reactor. As soon as the system is `brought tothe desired pressure of operation, e.g., about 300 pounds per .square Yinch gage, the temperature ofthe initial heater and reheaters is .increased from about `800" F. to about 850-1000 F., e.g., about 92`0` 4F., and the system is thus gradually brought to on-stream operating conditions.

By heating up .with flue gas, rather than Ystart-up gas, desorbed carbon oxides are removed, thereby avoiding contamination of start-up .gas and the necessity for repeatedpurging. By using nflue gas with asmall amount of oxygen, .e.g., 20.3 mol percent, for drying and establishing .circulation andinitial ,preheating conditions,Y catalyst activity .is protected. By introducing asubStantially hydrogenfree .start-up vgas to displace flue gas, carbon dioxideinuegas is vnot reduced by hydrogen Vto vcarbon monoxide, whichpoisons platinum catalyst. The impor tance of .removing the vunexpectedly-large amounts of desorbed .carbon'oxides before starting up a platinumalumina catalyst hydroforming system has been demonplatinum-alumina catalyst during heating.

10 ponents from the catalyst is minimized. .By .initiating the introduction of ncharging stock at low temperature and low pressure, carbon formation isminimized. Any carbon deposited on the catalyst during initial generation of hydrogen is-limited to the 'catalystin the -lreactor oper- 5 ating at higher temperature, preferably the swing reactor. Hydrogen produced' in this reactor will be adequate to prevent appreciable carbon formation in the remaining reactors when charge .is introduced-thereto. In systems including a swing reactor operating at'the 0 higher temperature, the swing-reactor may subsequently be cut out of the system, the catalyst thereinmay be regenerated, if needed, andthe other reactors will remain on-stream without necessity'for regeneration for a longer period of time. Y

In a preferred embodiment of our invention, westart u pV a platinum-alumina catalyst l'laphtha hydroforming system, finclu'ding lead, intermediate, and tail reaction zones,'heating.zoncsprior to each o'fsaid reaction =zones, a gasfseparation zone after said reaction zones, .and

compression means and lines for' recyclingseparatedtgas,

bypurging said reaction zones -With tile gas containing aboutV 0.2 .to 0.5 mol `percent oxygen,.circulating.said .liue

40 water ytherefrom and separating vthe condensed water from said fluegas in said .gasseparationzone depressuring, introducing a substantially :hydrogen-free. startfup :gas into fsaid reaction zones whereby. the partial` pressure of carbon oxides in said reaction zones'ispreducedvbelow .5 at least. about l poundv per square inch absolute, vintroducing preheated naphtha .charge stock at atemperature in the range of mtl-.800 F1. into said reaction zones, one of which reaction zones Vis ata higher Ytemperature than the remaining reaction pzones, while the pressureof 0 the startfup gas .is in the rangeoi about 3 0 to 200pounds per 'square .inch gage, continuing ztheyintroduction of naphtha charge stock at va ltemperature Yin said 'range until .the partialv `pressure of rhydrogen rreaches at .least rabout l pounds per square inch gage, and then in- 5 Acreasing the .temperature .of vthe naphtha change ,to at leastabout .850 F. f Y The invention will be more clearly understood by reference to the `following :exampleareadzin conjunetionxwith the accompanying drawing which is a schematic tow 0 diagram .of anUltraforming-:systemin which ourestart-up .procedureis` particularly advantageous. p Y

.In normal operation of an Ultraforming system a naphtha charge such, Afor example, las the 1'50 to 360 F 5 fraction ofpa Mid-Continent virgin naphtha vis-introduced transfer 'line' 10 from which the preheated chargemay be bypassed vby line 5.121 :to the lproduct recovery -system during start-up procedure. 'In oir-stream operation ,tr-anso fer line 10 will ,discharge Ythrough lines v12 and .13 to reactor 14 along with `recycled .hydrogen :from .line 15 which ispreheated .in ,heater 16. Eluent from .reactor V14 passes through line .17, Vreheater :18, vand .transfer Aline F19 itoreactor 20. .Eluent'from reactorlfpassese'through 5 .line lpreheaterfzz, and transterlinezzatoztail reactor 24.

from source 7 by pump 8. `'through preheater 'r9 and `than the initial reactors.

It should be understood that more than three reheaterreactor stages may be employed in the system.

Eluent from the tail reactor ilows through lines 25 and 25e, heat exchanger 26 and cooler 27 to separator 28 from which hydroformed product is withdrawn through line 29 to a stabilizer and/ or conventional product recovery system. A part of the hydrogen yWithdrawn from the separator through line 39 may be vented through line 31 but usually about 1,000 to 10,000 cubic feet per barrel of charge is recycled through lines 32 and 15a by means of circulating compressor 33 to line 15.

Transfer lines 11a, 13a, 19a and 23a may be selectively connected to header 34 for discharging through line 35 to swing reactor 36, the eiiiuent from which passes through line 37 to header 38 and thence through line 17a to line 17, line 21a to line 21, or line 25a to line 25. During normal on-stream operation without the swing reactor the valves in lines 11, 11a, 13a, 13b, 17a, 17h, 19a, 19b, 21a, 2lb, 23a, 23b, 25a and 25h remain closed, and the valves in lines 12, 13, 17, 19, 21, 23 and 25 remain open.

The swing reactor may be substituted for the lead reactor by opening valves in lines 13a, 35, 37 and 17a and closing valves in lines 13 and 17. Alternatively, it may be substituted for intermediate reactor 20 by opening valves in lines 19a, 35, 37 and 21a and closing the valves in lines 19 and 21. The swing reactor may take the place of the tail reactor by opening valves in lines 23a, 35, 37 and 25a, and closing valves in lines 23 and 25. It will thus be seen that each of the reactors may be taken olf-stream for regeneration and replaced by the swing reactor and that, alternatively, the swing reactor may be connected to operate in parallel with any of the other on-stream reactors during periods when no regeneration is required.

In some Ultraforming systems the hydrogen-rich recycle gas and the naphtha charge are heated in the same preheater. In such systems the charge introduced by pump 8 may be introduced by lines 8a and Sb to line 15 just ahead of heat exchanger 26 during normal operation and may be introduced by line 8a and line 8c to the line entering separator 28 prior to start-up.

Each of the reactors is provided With a refractory lining of low iron content, and metal surfaces may preferably be aluminized. They may each contain about the same amount of catalyst although, if desired, the subsequent reactors may contain somewhat more catalyst The catalyst may be of any known type of platinum-alumina catalyst. It may be prepared by compositing a platinum chloride with an alumina support as described, for example, in U. S. Patent 2,659,701, and it preferably contains about .1 to .8 weightv percent of platinum.

The on-stream pressure is usually below about 400 pounds per square inch gage, i.e. in the range of 200 to 350 pounds per square inch gage. The inlet temperatures to each reactor are usually in the range of about 850 to l000 F., c g. about 920 F., and may be approximately the same for each reactor although it is sometimes desirable to employ somewhat lower inlet temperature to the initial reactor than to the remaining reactors. The overall weight space velocity may be in the range of about 0.5 to pounds of naphtha per pound of catalyst per hour. -There is, of course, a pressure drop .in the system so that the lead reactor may operate at about 25 to 100 pounds per square inch higher pressure than the tail reactor.

Prior to regeneration hot hydrogen-rich gas for stripping hydrocarbons from catalyst in a blocked-out reactor maybe -introduced by line 41 to manifold line 42 and thence through one of lines 13b, 19h, 23b, or 35b to the selected reactor. Also, hydrogen-rich gas may be introduced from line to manifold line 39 via line 40.

For elfecting purging and regeneration of the catalyst in'any bed, purge gases and regeneration gasesmay be introduced through manifold line 39 and a selected one of lines 17h, 2lb, 2511 and 37b. Such purge and regeneration gases may be selectively Withdrawn through lines 13b, 1911, 23h and 35b to manifold line 42 from which gases may be 4vented or ared through line 43. Purge and regeneration gases from manifold line 39 may be introduced to the inlet of circulating compressor 33 by lines 44 and 15a.

Flue gas from source 45, which typically contains about 9 to l2 percent carbon oxides, about 18 to 14 percent water, and about 73 to 74 percent nitrogen, may be introduced to the system via regeneration facilities 46, which may contain compression facilities, furnaces, heat exchangers, gas puriers and the like, and valved line 47 to manifold line 39, and also to compressor 33 via lines 44 and 15a, when it is desired to introduce ilue gas into the system for startup, purging, and/or regeneration. Since carbon monoxide poisons platinum-alumina catalyst, particularly at high partial pressure, flue gas should also contain a small amount of oxygen (less than 1 mol percent) so that carbon oxides are present in the forni of carbon dioxide. This is particularly important when contacting platinum-alumina catalyst with flue gas at high pressure, but is less important at pressures not substantially above atmospheric, for instance, during a ue gas purge at atmospheric pressure. Air may also be intro` duced from source 4S via regeneration facilities 46 for effecting regeneration and/or regeneration-reiuvenation ofthe catalyst.

in starting up this system in accordance with our invention the entire system is, of course, rst checked for mechanical defects, cleaned out, dried, and the reactors are charged with catalyet. A layer of alumina balls is preferably placed on top of each of the catalyst beds to prevent swirling of the catalyst pellets which might otherwise lead to abrasion and production of catalyst iines. Flue gas from source 45 is next introduced into the system and passed via 46, 47, 39, 44, and 15a to the inlet of circulating compressor 33; and it is thereafter passed through all of the heaters, transfer lines and reactors, the flue gas being at this time purged fromthe system through line 31. The temperature of the flue-gas during this purge may be below about 200 F., and the pressure is preferably about atmospheric, i.e., about 5 to l0 pounds per square inch gage.

After the initial flue gas purge, the system is pressured with flue gas to at least the desired operating pressure, e.g., to about 300 pounds per square inch gage, to check for possible leaks.

After pressure testing, the valve in line 31 is preferably then set to hold back the full operating pressure, e.g., 300 pounds per square inch gage, and the reaction side is isolated from the regeneration side by closing valves in lines 47 and 50. The ue gas in the reaction side is continuously circulated through all the reactors, in series or parallel, preferably series, by means of compressor 33. For the preferred series ilow the ue gas after leaving compressor 33 would now via 15, 16, 13, 14, 17, 18, 19, 20, 21, 22, 23, 24, 25, 25e, 26,27, 28, 30, 32 and back to compressor 33 via 15a. 'I'he swing reactor is preferably connected in parallel With the tail reactor at this time. Thus, ue gas leaving furnace 22 would flow in part through line 23 to tail reactor 24, `and in part through line 23a, 34, 35 to swing reactor 36. Flue gas leaving swing reactor 36 would pass via lines 37, 38, and 25a so as to again join the other portion of ue gas leaving tail reactor 24 in line 25.

The temperature of the circulating flue gas is increased by gradually tiring furnaces 16, 18 and 22. Full operating pressure during this circulation step is preferred to assure sufficient circulation through furnaces and thereby avoid overheated tubes. The hot circulating gas etects drying of the catalyst and any further drying of the reactor limngs that may be required. At the same time carbon oxides adsorbed on the platinum-alumina catalyst Vmol percent water.

and linings .aredesorbed into thegcirculatug gases. To avoid fa very high ,concentration of l.carbon oxides in the circulating .flue ,gas resulting from desorption of .carbon oxides, llue gas .may he added to the circulating -ue .gas from lsou1ce.45 via 46 47, 39, 'and 44. At the same-.timefcirculating ue gas vmay be bled `from .the systemfviaflinel, or, alternatively, via 35b and 42. to vent 43.

'Water isfseparatedirom thecirculating Vgases =by condensation incooler27-.andlseparated in separator 28 from whichit vmay ;b.e withdrawn through line 53, the valve inlnei29 .beiugclosed at .this time. lHeating up of the reactionsystem-should be :sucinetly slow up to :about 300" tF.so thatthe gas -will Ynot zpick up more than 5 After .this initial drying step the temperature may be increased more rapidly to at least `about 50,0? F `at whichY temperature substantially all physically-.absorbed .and chemi-absorbed carbon oxides are-desorbed. At .this v.pointthe circulating flue gas may be vented from the reaction side and displaced with Ystart-up gas, egg., methane, Vsaid start-up gas .being substantiallyhydrogen-free, i.e., containing no more hydrogen than that,equivalent to a hydrogen partial pressure inithefreactionsystem of v.about 1 p.s.i.a. Of course, if

the partial'pressureof carbon oxides has already been reduced belowfabout v1 p.s.i.a., higher hydrogen partial pressures, eg.; above .about 45 `p.s.i..a., are feasible. Alternatively,..and,preferably, the circulating lluegas may be further heated to a temperature in the range of vabout 60G-800 F., preferably about 700-800 F., and then replaced withthe start-up gas. In eithercase circulation ofgases through thereactionwside should be continued at a temperature in the range of :about 60G-800 F., preferably about 7.00-800 F., until no further appreciable amounts of water-are removed from gas separation zone 28vialine 5.3. l Y

Y AWithssubstantially all adsorbed carbon oxides removed by the circulating flueL gas, the ue gas is removed ,from the systemY by depressuring vialines 31 or 43 and, prefferablyybyjevacuation (facilities not shown). Start-up gas, .whichmay .further displace flue gas, is introduced to apressure'in excess of .about 30 pounds per square inch gage, fromsource V52 .via linef4 9, 39, 44, and 15a to lthe inlet ofcompressor. 33. Partial pressure of carbonoxides in the reaction system after introduction of start-up gas should 'bebelow about lrpound per'square :inch absolute, aflevel .which ,euperien'cehas .shown to cause .negligible deactivation; AIf the lreactionside has not already been heatedto a temperature of about 600-800 F., preferably about y700-800 F., the Vstart-upgas should be circulated throughthesystem to; raise the temperature to that level by means of;furnaces 16, 18, and 22. This circulation should, of course,.continue until no nfurther appreciable amounts of water-.areremovedfrom:gas separation zone 28-via line v53. ,f

With essentially undilutedstart-up .gas being recycled through the reaction Hside, thevalve in line 53 is closed and the valve in line .29 is opened. Naphtha preheater 9 is tired and charge naphtha preheated to about 800 F. isintroducedthrough by-pass line 11 (the valves in lines .11a and12 being closed) directly to product ethuent line V25::rom which it passes through exchanger 26, cooler 27,-.separator-28, -and line 29 to the product recovery sysvtem'forestablishingoperating conditions therein. If a single-preheater isV employed, naphtha charge in introduced atcthisV time to the separatorfthrough lines 8a and 8c, the valve in line Str-being closed. Naphtha should not .of course, beiintroducedfwhile due-.gas is being circulated. therwise,.naphtha vapors may be entrained by the flue v.gas in separator 28 yand carried to the reaction system, thereby/.resulting inproduction of hydrogen, reductionof carbon vdioxide .to carbon monoxide, and Ypoisoning of thecatalyst.

The pressure of the, hot circulating start-up gasif it is `riotalready below about200 pounds per square inch gage,

shouldnest he reduccdtoas .lowa pressure asl iseasible; preferably to a pressure inthe .range of aboutnto :1.50 ipoundspersquare inch'gage, by .adjustment of the .-pres sure cont-rol yvalve in line 31. The lower limit ongpressure is usually dictated by the minimum suction pressure of the recycle .gas compressor 133. A jlow start-:up 4:gas pressure is desired from a process standpoint so Vthat the partial :pressure of hydrogen will increase as rapidly ,as possibiafonce naphtha is charged.

The catalyst in'one of the reactorsis preferably heated to about 750-800 F.,iand the naphtha which isintroduced into this reactor is preheated to about the same temperatureso thatthe generation-,of hydrogen will largely :be accomplished in arsingle vessel. When .the-swing reactor is employed for yhydrogen generation .and vall catalyst bedshave been heated to about TO0-750 F. by .circulating vstart-up gas, the valvein ,line .23 may be closed, .the fvalvein line'23a openedand ,the temperature Y of .heater 22 increased toprovide-a transfer line temperatureof .about 800 F. so that thesWing reactor isithus heated to about 7 50-800" F Next, withrthe pressurcfof the circulatingstart-up gas below about l.200 pounds @per square inch gage andpreferablyqin the range of 30 to 150 poundsper square inchgage, the valves in lines 11a and .23 .fa-reopened, While valvesi-n lines 11, 12 and 13a are closed. Naphtha vapors preheated to approximately-800 F. are introducedqdirectly through lines 34 and 35 to the .swing reactor while recycled gases ,are continuously circulated through heater 16,.reactor 14, heater 18,;reactor 20, `heater 422 and reactors 24 and 36 (in parallel), the .temperature of these heaters at this time being increased to bring the temperature of the lead and intermediate reactors .up to about SGU-.850 F.

The hydrogen .generated by dehydrogenation in .the swing reactor will quickly increase the hydrogen concentrationof. the recycled gas to about .80fpercentor more andthe pressure in the-system may be increased by the generated hydrogen to the desired operating pressure. As soon as the hydrogenpartial` pressure builds up to at least about. pounds per squarev inch gage, the valve in line 12'is opened Aand the valve in line v11a is closed so that the charge preheated to about.800 F.. is now cutiuto vthe hot `.circulating.hydrogen stream. The .transfer line temperatures of heaters 9, `1.6, y18 and 22 may now be increased to the desired level, e.g. about 900-f950 F. Thus, carbon .deposition on the lead, intermediate, .andtail-reactorsissubstantially ,avoided by starting up with hydroygen produced in theswing reactor. If and when regeneration of the catalyst in the swing Yreactor is desired, the valvein line 35 is closed and sucienthothydrogen is .available for stripping hydrocarbons therefrom valines 41,42 and 35b, after which thc valves in lineszi41, 37, and 25a are closed.

AInstead of employingthe swing reactor .for initially generating hydrogen, thelead reactor, the tail reactoror, in.fact,.any of the on-streamereactorsl may be .preheated to a temperature -of about 750800 F. by raising the temperature Aof the appropriate heater, and the .system may bev brought on-strearn by introducing charging stock at about 7.00" F. to all of thereactors-except the preheated reactor for which the starbup inlet temperatureV is preferably about '750-800" F. With this method of/.start-.uln lineV 11a'is not required and the preheatednaphtha may be introducedto. line 13 .andreactor 14 viatrans,fer lines 410 andlZ when a separate preheater 9 is employed .or

may beu introduced via lines 8a, 8b, and 15 when hydrogen and naphtha are preheated in the same heating :9.0i1.

. `In this embodiment lone reactor, eg., thesWing-rehydrogen partial pressure has reachedabout Y150.p.ounds pervsquare inch. gage. When .thecirculating start-upgas .has thus been essentiallyreplaced by hydrogen .and .the

partial pressure of thevhydrogen has -been .increased-.toast 9 least about 150 pounds per square inch gage, all transfer line temperatures may be increased to the desired operating level and any further increase in pressure may be built up so that on-stream operating conditions are fully established. In this case, the reactor which has operated at the initially high temperature and in which most of the carbon deposition has occurred may be cut out of the system for regeneration and replaced by the reactor which was initially blocked out.

The method of effecting catalyst regeneration will be described as applied to the swing reactor 36 but it should be understood that the same procedure may be employed for any one of the other reactors when it is blocked out. When the charge inlet valve in line 35 is closed and while the valve in line 37 remains open, hot hydrogen-rich recycle gas is introduced by line 41 to manifold line 42 and thence through line 35b to strip out any hydrocarbons that may remain in reactor 36, this stripped material being discharged through lines 37, 3S, and 25a to line 25. Next the valve in lines 41, 37, and 25a are closed and reactor 36 is depressured by opening the valve in line 43. Next, the reactor is purged to eliminate hydrogen therefrom by introducing flue gas from source 45 via 46, 47, 39 and 37b, the purge gases being vented through lines 35b, 42 and 43. The temperature of the catalyst bed is adjusted to about 650-800 F. preparatory to initiating regeneration by circulating llue gas, under approximately the same pressure as that employed in on-stream processes, i.e., about 300-350 pounds per square inch gage, and controlling llue gas temperature at its source 45 or in regeneration facilities 46. Next, controlled amounts of air are introduced from source 48 into the circulating ue gas stream at a rate to effect combustion of carbonaceous deposits without exceeding a combustion zone temperature of about l050 F. The hot ilue gas leaving reactor 36 at about this temperature is vented via lines 35b, 42, and 43 or is passed by lines 35h, 42 and 50 back to regeneration facilities 46 wherein the gas is scrubbed to eliminate most of the water formed by combustion of hydrocarbonaceous deposits and the net amount of llue gas production is vented from the system through line 51.

If rejuvenation is required (and it may not be required until the catalyst has been regenerated many times) the introduction of flue gas is stopped and the introduction of air is continued so that the catalyst is treated with a circulating air stream at a pressure of about 300 pounds per square inch gage and a temperature of about 950 F. or more for a period of about one-half hour to twelve hours or more depending upon the extent of rejuvenation required. For rejuvenation suflicient air must be added so that the partial pressure of the oxygen is at least 0.4 atmosphere.

After the regeneration (or after rejuvenation if rejuvenation has been elected) the introduction of air is stopped, and ue gas is introduced from source 45 to purge oxygen from the swing reactor and from the regeneration system via line 51 until the oxygen content is reduced below about l mol percent. After this highpressure flue gas purge, the system is depressured by slowly opening the valve in line 43 and ilue gas is introduced from source 45 to purge substantially all remaining oxygen from the swing reactor and from the regeneration system. After the flue gas purge, the system may then, preferably, be evacuated and/ or purged with startup gas, e.g., methane from source 52, following which hydrogen-rich recycle gas is introduced through lines 40, 39 and 37b. The methane purge serves to separate carbon dioxide in flue gas from hydrogen in recycle gas, and thus avoids formation of carbon monoxide. Preferably, to suppress heat front formation, the methane may contain about 0.5 mol percent hydrogen sulfide. After introduction of recycle gas, the valve in line 37b is closed and the reactor is pressured with bot hydrogen introduced by lines 4l, 42 and 35b. When the reactor is thus brought to desired operating pressure, the valves in lines 41 and anantie i0 35h are closed, and the reactor may be placed oil-stream by opening valves in lines 3S and 37.

While our invention has been described with respect to a particular Ultraforming system, it should be understood that it is applicable to other types of platinum catalyst hydroforming systems, regenerative or unregenerative, in which the platinum catalyst may contain adsorbed carbon oxides. Various alternative arrangements and operating conditions will be apparent in the above description to those skilled in the art.

Having thus described our invention, We claim:

l. The method of starting up a naphtha hydroforming zone containing platinum-alumina catalyst with adsorbed carbon oxides at temperatures below about 500 F. which method comprises purging said zone with flue gas, said ilus gas containing carbon dioxide and about 0.1 to l mol percent oxygen, circulating said flue gas through said zone at a pressure in excess of about 30 pounds per square inch gage while simultaneously heating the circulating flue gas to a temperature above about 500 F., depressuring, introducing at a temperature above about 500 F. and a pressure in excess of about 30 pounds per square inch gage a substantially Vhydrogen-free startup gas, said startup gas being selected from the group consisting of nitrogen, normally-gaseous hydrocarbons, and mixtures thereof, into said Zone, whereby the partial pressure of carbon oxides in said zone is reduced below about l pound per square inch absolute, circulating said Startup gas through said zone until said zone is at a temperature in the range of about 600 to 800 F., introducing into said zone while the pressure of the startup gas therein is in the range of about 30 to 200 pounds per square inch gage preheated naphtha charge stock at a temperature in the range of 650-800 F., whereby hydrogen is produced, continuing the introduction of napbtha charge stock at a temperature in said range until the partial pressure of hydrogen in said zone reaches at least about pounds per square inch gage, and then increasing the temperature of the naphtha charge to at least about 850 F.

2. The method of claim l which includes the step of simultaneously increasing the pressure in said hydroforming zone while the hydrogen concentration therein is being increased to at least 150 pounds per square inch gage.

3. The method of claim l wherein said depressuring includes the step of evacuating said hydroforming zone prior to introducing said startup gas.

4. In a regenerative naphtha hydroforming system including a reaction side with lead, intermediate, and tail reaction zones containing platinum-alumina catalyst, heating zones prior to each of said reaction zones, a gas separation zone after said reaction zones, and compression means and lines for recycling separated gas, and including a regeneration side with a source of liue gas and a source of oxygen for regeneration, and Valved lines for connecting said regeneration side with said reaction side, the method of start-up from temperatures below about 500 F. which comprises purging said reaction side with ue gas introduced from said regeneration side, circulating ue gas containing about 0.1 to l mol percent oxygen, whereby said oxygen concentration assures that the carbon oxides in said flue gas are carbon dioxide, through said reaction side at a pressure in excess of about 30 pounds per square inch gage while simultaneously heating the circulating tlue gas to a temperature above about 500 F., depressuring said reaction side, introducing at a temperature above about 500 F. and a pressure in excess of about 30 pounds per square inch gage a substantially hydrogen-free startup gas, said startup gas being selected from the group consisting of nitrogen, normallygaseous hydrocarbons, and mixtures thereof, into said reaction side whereby the partial pressure of carbon oxides in said reaction side is reduced below about l pound per square inch absolute, circulating said startup gas through said reaction side until said reaction side is at a temperature in the range of about 700 to 800 F., introducing preheated naphtha charge stock at atemperature in the range of 700 to y800" F. :into at least one of said reaction z ones while the pressure of .the .start-,up gas is in the range of. about 30 to 200 pounds per square inch gage, whereby lhydrogen is produced, continuing the introduction of naphtha charge vstock at a temperature in said range until thepartial pressure of .hydrogen in said reaction side reaches at least about 150 pounds per square inch gage, -andthen increasing the temperature of the naphtha charge to at least about .850 IF.

5. The-method lof claim 4 wherein depressuring said reaction side includes the ,stepof evacuating said reactionside prior to introducing'said startup gas.

i6. The methodof claim 4 which includes the step, while `circulating-said start-up gas through said reaction side at a temperature `in the range of about 700-800 F., ofremoving from saidstart-up gas any water removed 'from'said reactionside by said .start-up gas, prior to introducingnaphtha into said reaction side.

7, The method of claim 4which includes the step of initially introducing preheated naphtha charge stock into Ya reaction zone which is. at a higher temperature than the remaining reaction zones.

8. Themethod of claim 4 in which the reaction side contains a :swing reaction zone, which method includes Vthe steps Yofvinitially passing preheatednaphtha charge stock Ionly through the swing reaction zone so as to .provide vthe hydrogen build-up in the system and subsequently-blocking out said swing reactionzone and regenerating catalyst therein.

` 9. The method of claim 4 wherein, during the step of circulating ue gas through said reaction side =while simultaneouslyheating the circulating flue gas to a temperature Vabove V.about 500 F., additional ue gas .is introduced from said regeneration side to said reaction side and a portion of said circulating flue gas is vented from said reaction side, whereby the buildup of carbon oxides in said circulating flue gas, caused by desorption of carbon oxides from `s aid catalyst, is minimized.

10. In a platinum-alumina catalyst naphtha hydroforming system including lead, intermediate, and tail reactionzones, heating zones prior to each of said reaction zones, a gasseparation zone after said reaction zones, and compression means and lines for recycling .separated gas, .the -niethod v.of start-uptrend .temperatunes below about 50.09 E. .which vco1iziprises purging said reaction zones withflue gas-cont ingabout 0.2 to 0,5 :mol per' centi oxygen, e whereby oxygen concentration .assures that thec-arbon .Oxides in saidfie gas are .carbon dioxide, circulating lsaid uelgasthrough .said reaction zones at a pressure in excessof about 30 pounds `per square inch vgage while simultaneously Yheating the circulating Yflue gas to a temperature in the rangeoffabout T100-.800 .F continuingcirculation of .saidue lgas through said lre'- action zones at atemperature Ainsaid range while simultaneously cooling said ue gas leavingsaid tail reaction zone to condense .water therefrom and separating .the condensedwater from .said flue gas in .saidgas separation zone, depressuring vintroducing at va't.ernperature-in range of about 709-800" F. and a pressurefin excesslof about 30 pounds per square `inch gage asubstantially hydrogen-.free start-nngas, said startup'k gasbeing. selected from the group consistingrof nitrogen, Anormal.1y-gaseous hydrocarbons, vand mixtures thereof, `intozsaid .reaction zones whereby the partial pressure of carhonozddes .in said .reaction zones is .reduced below at leastaboutj pound per square :inch absolute, introducing preheated -naphtha charge .stock :at a temperatureinthe range of 7004500" F, intosaid reaction zones, one of which reaction v.zones -isata higher temperature-than the remaining reaction zones, -whilethe pressure of thestart-up .gas is in the mangent about30 to 200pounds per squareineh gage,vwhereby hydrogenis produced, `continuing the .introduction of 4naphtha :chargestockat a temperaturerin .said rangeuntil the partial pressure of-hydrogen is increased to at leastfabout 150 pounds per squarev inch gage, and then increasing-the .temperatur-eef the naphtha charge to at least about .850 F.

kReifen-en lees Cited .in the file o f this patent `UNITED .STATES PATENTS 

1. THE METHOD OF STARTING UP A NAPHTHA HYDROFORMING ZONE CONTAINING PLATINUM-ALUMINA CATALYST WITH ADSORBED CARBON OXIDES AT TEMPERATURES BELOW ABOUT 500*F. WHICH METHOD COMPRISES PURGING SAID ZONE WITH FLUE GAS, SAID FLUE GAS CONTAINING CARBON DIOXIDE AND ABOUT 0.1 TO 1 MOL PERCENT OXYGEN, CIRCULATING SAID FLUE GAS THROUGH SAID ZONE AT A PRESSURE IN EXCESS OF ABOUT 30 POUNDS PER SQUARE INCH GAGE WHILE SIMULTANEOUSLY HEATING THE CIRCULATING FLUE GAS TO A TEMPERATURE ABOVE ABOUT 500*F., DEPRESSURING, INTRODUCING AT A TEMPERATURE ABOVE ABOUT 500*F. AND A PRESSURE IN EXCESS OF ABOUT 30 POUNDS PER SQUARE INCH GAGE A SUBSTANTIALLY HYDROGEN-FREE STARTUP GAS, SAID STARTUP GAS BEING SELECTED FORM THE GROUP CONSISTING OF NITROGEN, NORMALLY-GASEOUS HYDROCARBONS, AND MIXTURES THEREOF, INTO ZONE, WHEREBY THE PARTIAL PRESSURE OF CARBON OXIDES IN SAID ZONE IS REDUCED BELOW ABOUT 1 POUND PER SQUARE INCH ABSOLUTE, CIRCULATING SAID STARTUP GAS THROUGH SAID ZONE-UNTIL SAID ZONE IS AT A TEMPERATURE IN THE RANGE OF ABOUT 600 TO 800*F., INTRODUCING INTO SAID ZONE WHILE THE PRESSURE OF THE STARTUP GAS THEREIN IS IN THE RANGE OF ABOUT 30 TO 200 POUNDS PER SQUARE INCH GAGE PREHEATED NAPHTHA CHARGE STOCK AT A TEMPERATURE IN THE RANGE OF 650-800*F., WHEREBY HYDROGEN IS PRODUCED, CONTINUING THE INTRODUCTION OF NAPHTHA CHARGE STOCK AT A TEMPERATURE IN SAID RANGE UNTIL THE PARTIAL PRESSURE OF HYDROGEN IN SAID ZONE REACHES AT LEAST ABOUT 150 POUNDS PER SQUARE INCH GAGE, AND THEN INCREASING THE TEMPERATURE OF THE NAPHTHA CHARGE TO AT LEAST ABOUT 850*F. 